WO2013135986A1 - Optimized method for recycling bio-oils into hydrocarbon fuels - Google Patents

Abstract

The present invention relates to a method for producing hydrocarbon materials from a feedstock including at least one non-pretreated bio-oil, said feedstock being supplied in at least a first step of hydroreforming in the presence of hydrogen and a hydroreforming catalyst including at least one transition metal selected from among the elements from groups 3 through 12 of the periodic table, and at least one support selected from among activated carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide, and transition-metal aluminates, taken alone or in a mixture, so as to obtain at least one liquid effluent including at least one aqueous phase and at least one organic phase, and a second step of hydrotreating and/or hydrocracking at least part of the organic phase of the effluent from the hydroreforming step so as to obtain a hydrocarbon effluent, wherein at least part of the organic phase of the effluent from the hydroreforming step is recycled in the hydroreforming step, mixed together with said feedstock including at least one bio-oil, at a recycling rate equal to the ratio of the mass flow of said organic phase to the mass flow of the non-pretreated bio-oil, which is between 0.05 and 2, and wherein the hydrocarbon effluent from the hydrotreating and/or hydrocracking step is not recycled in said hydroreforming step.

Description

 OPTIMIZED PROCESS FOR THE VALORISATION OF BIO-OILS TO HYDROCARBON FUELS

Field of the invention

 The present invention relates to a process for producing hydrocarbon products, in particular chemicals and / or hydrocarbon products that can be incorporated in the pool of liquid fuels called biofuels, from liquids derived from biomass and in particular from liquefied biomass pyrolysis of biomass, also called bio-oil. More specifically, the invention relates to a process for recovering bio-oils into hydrocarbon products that can be incorporated into the pool of liquid fuels for engines intended for air, sea or land transport applications, by means of a catalytic process in two. steps, said method comprising a hydroreforming step, followed by a hydrotreatment and / or hydrocracking and / or hydroconversion step.

 State of the art

 Bio-oils are liquid products obtained by thermochemical liquefaction of cellulosic biomass. Thermochemical processes generally convert biomass into liquid, gaseous and solid products. Among these methods, so-called rapid pyrolysis methods tend to maximize the liquid yield. During rapid pyrolysis, the temperature of the biomass, optionally finely divided, is rapidly raised to values of greater than about 300 ° C. and preferably of between 300 and 900 ° C. and the liquid products are condensed in the form of biocompounds. oil. Ringer et al. (Large-Scale Pyrolysis OH Production: A Technology Assessment and Economics Analysis, M. Ringer, V. Putsche, and J. Scahill, NREL Technical Report NREL / TP-510-37779, November 2006) studied the different technologies implemented. for rapid pyrolysis of biomass on a large scale. They include bubbling fluid beds, circulating fluidized beds, ablative pyrolysis reactors, vacuum pyrolysis and rotary cone pyrolysis.

 At first glance, bio-oils derived from the rapid pyrolysis of biomass could in principle provide low-cost renewable liquid fuels; indeed, their use as fuel for boilers, stationary gas turbines and diesels has been described. In addition, the rapid pyrolysis has been developed industrially on a relatively large scale, of the order of several hundred tons per day. However, this technology has not yet attracted much commercial interest.

 Indeed, a bio-oil and in particular a bio-oil resulting from the rapid pyrolysis of biomass is a complex mixture of highly oxygenated polar hydrocarbon products obtained from the breaking of biopolymers, and having physical and chemical properties limiting its direct use as biofuel.

Pyrolysis bio-oils have undesirable properties such as: (1) corrosivity due to their high water and organic acid content; (2) a low specific calorific value due to the high oxygen content, typically of the order of 40% by weight; (3) chemical instability due to the abundance of reactive functional groups such as the carbonyl group, which can cause polymerization during storage and hence phase separation; (4) relatively high viscosity and phase separation tendency under shear conditions high, in an injector for example; (5) an incompatibility with conventional hydrocarbon fuels because of their insolubility therein; (6) solid particles of coal resulting from pyrolysis, which will always be present in unfiltered bio-oils to a greater or lesser degree, may cause plugging of injectors and pipes. All these aspects combine to make the handling, transportation, storage and implementation of bio-oils difficult and costly, thus making their integration into current heat and energy production systems and technologies problematic.

 The economic viability of the production of bio-oils for energy applications thus depends on the discovery of suitable methods for their development into a high-quality liquid fuel, at a sufficiently low cost. Indeed, considerable efforts have been made in recent years to seek practical technologies that can overcome all or some of the disadvantages mentioned above.

 One of the proposed approaches is to esterify and acetylate the bio-oils with alcohols such as ethanol and butanol as for example in the patent application EP 0 718 392. The reaction products nevertheless always have a high acidity and a high water content, while the increase in the specific heating value is low. In addition, the products themselves tend to be chemically unstable and reactive.

 Another approach is to emulsify the bio-oil diesel fuels using suitable surfactants such as in US Patent 5,820,640. Although the problem of compatibility with industrial hydrocarbon fuels is thus solved, new problems such as the stability of the emulsion occur, while those related to corrosivity and chemical stability remain.

 In yet another approach, the non-condensed bio-oil vapor resulting from the rapid catalytic pyrolysis of solid biomass is partially deoxygenated by transformation in the presence of zeolitic catalysts in order to directly produce low molecular weight aromatics such as BTX ( benzene, toluene, xylene) as for example in US Pat. No. 4,308,41 1. However, the zeolitic catalysts are acidic and the oxygen is largely removed by dehydration to give water. Consequently, because of the inherent molar deficiency of hydrogen with respect to oxygen and carbon in the bio-oil, the organic liquid fraction yields are relatively low and a very large formation of coke is observed, which increases the technological difficulty and affects economic performance.

An indirect approach is to gasify the bio-oil (and / or carbonization by-product) into synthesis gas, then to perform a Fischer-Tropsch synthesis of long-chain hydrocarbons or olefins derived from the synthesis gas in the context of a process known as BTL (Biomass To Liquid) (for example the Bioliq process described by Henrich et al (Cost estimate for biosynfuel production via biosyncrude gasification, E. Henrich, N. Dahmen and E. Dinjus, Biofuels, Bioprod Bioref 3: 28-41 (2009)) Overall hydrocarbon yields from biomass are nevertheless relatively low and investment costs are high. In addition, the minimum volumes from which BTL processes are expected to be economical are large relative to the typical local availability of biomass, which requires complex logistics for biomass supply, with significant transportation costs.

 Over the last twenty years, the approach of direct hydrotreating of bio-oil to stable hydrocarbons or oxygenated liquids has been intensively studied. Elliott has published a detailed study of these many historical developments, including work with model compounds known to be present in bio-oils (Historical Developments in BioOils Hydroprocessing, DC Elliott, Energy & Fuels 2007, 21, 1792-1815). ).

 One of the major impediments to direct catalytic hydrotreatment of bio-oils is their propensity to polymerize when subjected to temperatures above about 100 ° C, which results in the formation of solids at elevated temperatures. greater than about 140Ό, with consequences such as the clogging of charge lines, furnaces and reactors as well as the rapid deactivation of the catalyst, or even clogging of the reactor inlet.

 These difficulties can be partially overcome by hydrotreating only the thermally resistant portion of the bio-oil. Thus, by adding water to the bio-oil, it can be separated into an aqueous phase and into a viscous phase of higher density, generally between 20 and 30%, called pyrolytic lignin because amply derived from the ligneous fraction. the pyrolysis load of the biomass. Since the pyrolytic lignin is rich in phenolic material, it has a much higher thermal stability than that of the carbohydrate-derived portion of the bio-oil, and is therefore easier to catalytically hydrotreat without formation of solids. This is the approach followed by Piskorz et al. (Conversion of Lignins to Hydrocarbon Fuels, Piskorz J., Majerski P., Radlein D., Scott DS, Energy & Fuels, Vol 3, 723-726, 1989) and, more recently, Marker and Petri in US Pat. 7578927. This method has the disadvantage of implementing a step of separating the bio-oil upstream of the catalytic hydrotreating step and the obligation to treat the two resulting fractions separately. The resulting process is therefore complex and its implementation leads to losses in mass and energetic efficiency, as well as in operating cost.

 The hydrotreatment of the complete bio-oil, that is to say having not undergone any prior separation step was considered. US Pat. No. 4,795,841 proposes a process for the hydrotreatment of the complete bio-oil in two stages, the bio-oil being used alone, ie not mixed with other streams, in the first step of which the thermal stability The overall bio-oil is improved by catalytic hydrogenation at a temperature of about 250 ° C.

However, this two-stage hydrotreatment process results in the formation of solids and the rapid deactivation of the catalyst. Hydroprocessing of bio-oils is also difficult because of the formation of multiple organic phases, in addition to the aqueous phase. Moreover, the kinetics of the hydrogenation is slow, especially at the relatively low temperatures required for the thermal stabilization of bio-oils. Hydrogen consumption is also high.

 The patent application US-2009 / 0,253,948 discloses a process for converting pyrolysis oil derived from biomass into a liquid fuel by means of a two-stage deoxygenation of the pyrolysis oil and separation of the products, in which the final hydrocarbon product as well as the fractions resulting from the final distillation and boiling at boiling points corresponding to the fractions that are valuable in the gasoline, kerosene and diesel fuel pools can be recycled. However, it does not teach the high degree of dispersion or solubilization of the pyrolysis oil in a hydrocarbon medium, with the advantages of a sharp increase in reaction rates and catalyst life.

 Nevertheless, it remains to find an efficient hydrotreatment process. The difficulties encountered have their origin in the rapid thermal polymerization of the bio-oil, which leads to line clogging and reaction equipment as well as a rapid deactivation of the catalyst. In other words, at the temperatures typically required for the hydrotreatment of bio-oils, the polymerization reactions are substantially faster than the hydrotreatment reactions, which ultimately results in the formation of coke.

 There is therefore a need for an improved bio-oil hydrotreatment process that minimizes solids formation, catalyst deactivation and hydrogen consumption, and maximizes the deoxygenated oil yield produced. It is also desirable to maximize the fraction of the hydrocarbon product whose boiling point is in the boiling range of useful fuels such as gasoline, kerosene or diesel.

 Summary of the invention

 The present invention relates to a method for producing hydrocarbon products from a feedstock comprising at least one non-pretreated bio-oil, said feedstock being introduced in at least the following stages:

 a first hydroreforming step in the presence of hydrogen and a hydroreforming catalyst comprising at least one transition metal chosen from the elements from groups 3 to 12 of the periodic table and at least one support chosen from activated carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, taken alone or as a mixture, to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase,

 a second step of hydrotreating and / or hydrocracking at least a portion of the organic phase of the effluent from the first hydroreforming step to obtain a hydrocarbon effluent,

wherein at least a portion of the organic phase of the effluent from the first hydroreforming step is recycled to said first hydroreforming step, in admixture with said feed comprising at least one bio-oil, with a recycle ratio equal to the mass flow rate ratio of said organic phase on the mass flow rate of the non-pretreated bio-oil of between 0.05 and 2 and in which the hydrocarbon effluent from of the hydrotreatment and / or hydrocracking step is not recycled in said first hydroreforming step.

 An advantage of the present invention is to provide a two-stage process for the production of a stable hydrocarbon effluent that can be incorporated in the fuel pool, from a feedstock comprising at least one non-pretreated bio-oil, comprising a lower oxygen content. at 5% and preferably less than 1% by weight and a water content of less than 0.1% by weight relative to the total mass of said hydrocarbon effluent, without significant formation of coke or bio-oil polymers, with low speeds deactivation of the hydroreforming catalyst, substantially improved conversion rates and reduced hydrogen consumption.

 Another advantage of the present invention is to treat a feedstock comprising at least one bio-oil that has not undergone any pre-treatment step. The operating cost of the process according to the invention is therefore reduced.

Another advantage of the present invention is to provide a process having a great flexibility with respect to the nature of the liquids derived from the biomass that can be used as feedstock comprising at least one non-pretreated bio-oil, in particular liquids resulting from the pyrolysis of biomass, which may have broad ranges of viscosity, water content and degree of polymerization, depending on the origin of the biomass and the pyrolysis process carried out.

 The inventive method disclosed herein satisfies all of these criteria.

 Detailed description of the invention.

 According to the invention, the bio-oil of the filler treated in the process according to the invention is not pretreated, that is to say that said bio-oil has not undergone any pretreatment step other than pyrolysis. In particular, said bio-oil has not undergone any pretreatment step intended to reduce its metal content, its solids content, its water content or its acid content.

 A bio-oil is a complex mixture of oxygenates, obtained from breaking biopolymers present in the original biomass. In the case of lignocellulosic biomass, the structures from these three main components, cellulose, hemicellulose and lignin, are well represented by the components of the bio-oil.

In particular, a bio-oil is a highly oxygenated polar hydrocarbon product which generally comprises at least 10% by weight of oxygen, preferably from 10 to 60% by weight, and preferably from 20 to 50% by weight of oxygen. relative to the total mass of said bio-oil. In general, the oxygenated compounds are alcohols, aldehydes, esters, ethers, organic acids and aromatic oxygen compounds. Part of the oxygen is present in the form of free water representing at least 5% by weight, preferably at least 10% by weight and preferably at least 20% by weight of the bio-oil. These properties make the bio-oil totally immiscible with hydrocarbons, even with aromatic hydrocarbons that usually contain very little or no oxygen.

 The bio-oils are preferably derived from biomass preferably selected from plants, herbs, trees, wood chips, seeds, fibers, seed husks, aquatic plants, hay and other sources of lignocellulosic materials, such as those from municipal waste, agro-food waste, forest waste, slaughter residues, agricultural and industrial waste (such as sugar cane bagasse, waste from oil palm cultivation, sawdust or straws). Bio-oils can also come from paper pulp and by-products of recycled or non-recycled paper or by-products from paper mills.

 The bio-oils are advantageously obtained by thermochemical liquefaction of the biomass, preferably by pyrolysis, and preferably by rapid, slow pyrolysis with or without a catalyst (in the presence of a catalyst, this is known as catalytic pyrolysis). Pyrolysis is a thermal decomposition in the absence of oxygen, with thermal cracking of the gas, liquid and solid feedstocks. A catalyst may advantageously be added to enhance the conversion during so-called catalytic pyrolysis. Catalytic pyrolysis generally gives a bio-oil having a lower oxygen content than the bio-oil obtained by thermal decomposition, but the yield of bio-oil is generally lower. The selectivity between the gas, the liquid and the solid is related to the reaction temperature and the residence time of the vapor. Biomass pyrolysis processes, including rapid pyrolysis, are well described in the literature such as, for example, A.V. Bridgewater, H. Hofbauer and S. van Loo, Thermal Biomass Conversion, CPL Press, 2009, 37-78.

 The slow pyrolysis is advantageously carried out at a temperature of between 350 and 450 Ό and preferably of the order of 400 ° C. and with a residence time of a few minutes to a few hours which favors the production of a solid product as well. called carbonization product or charcoal. In particular, slow pyrolysis generally allows the production of 35% by weight of gas, 30% by weight of liquid and 35% by weight of char. Gasification processes operating at very high temperatures, preferably above 800 ° C, promote the production of gas, and in particular more than 85% by weight of gas. The intermediate pyrolysis is advantageously carried out at a temperature generally of between 450 and 550 ° C. and with a residence time of the short vapor, preferably between 10 and 20 seconds, which favors the liquid yield. In particular, the intermediate pyrolysis generally allows the production of 30% by weight of gas, 50% by weight of liquid and 20% by weight of char.

The rapid pyrolysis is advantageously carried out at a temperature generally of between 450 and 550 ° C. and with a residence time of the very short vapor, preferably between 0.5 and 2 seconds, which maximizes the liquid yield: in particular, the rapid pyrolysis generally allows the production of 10-20% weight of gas, 60-75% weight of liquid and 10-20% weight of carbonization product. The highest liquid yields are therefore obtained in fast pyrolysis processes, with a liquid yield of up to 75% by weight.

 Preferably, the bio-oils used in the process according to the present invention are obtained by rapid pyrolysis of the biomass.

 The filler treated in the process according to the invention preferably comprises a bio-oil content of between 10 and 100% by weight relative to the total weight of said filler.

 The filler comprising at least one non-pretreated bio-oil used in the process according to the invention may also advantageously contain other compounds derived from biomass, said compounds being advantageously chosen from vegetable oils, algal or algal oils, fish oils, fats of vegetable or animal origin, and alcohols derived from the fermentation of biomass sugars, or mixtures of such fillers, whether pre-processed or not. Vegetable oils or oils derived from animal fats essentially comprise chemical structures of the triglyceride type which are also known to those skilled in the art as tri-ester of fatty acids as well as free fatty acids, the fatty chains of which contain a fatty acid. number of carbon atoms between 9 and 25.

 Said vegetable oils may advantageously be crude or refined, wholly or in part, and derived from plants selected from rapeseed, sunflower, soybean, palm, olive, coconut, copra, castor, cotton , peanut oils, flax and crambe and all oils from eg sunflower or rapeseed by genetic modification or hybridization, this list is not limiting. Said animal fats are advantageously chosen from bacon and fats composed of residues from the food industry or from the catering industries. Fried oils, various animal oils such as fish oils, tallow, lard can also be used.

Said feedstock comprising at least one bio-oil may also advantageously be treated in admixture with at least one hydrocarbon feedstock derived from petroleum and / or coal. The hydrocarbon feedstock derived from petroleum can advantageously be chosen from the direct distillation distillates under vacuum, the vacuum distillates resulting from a conversion process, such as those resulting from coking, fixed bed hydroconversion or hydrotreatment processes. ebullated bed heavy fractions, products from fluid catalytic cracking units, such as, for example, light catalytic cracked gasoil (LCO) of different origins, heavy catalytic cracked gasoil (HCO) of different origins and any fraction fluid catalytic cracked distillate product generally having a distillation range of about 150Ό to about 370Ό, aromatic extracts and paraffins obtained in the preparation of lubricating oils and solvent deasphalted oils, alone or in admixture. The hydrocarbon feedstock derived from coal may advantageously be chosen from products resulting from the liquefaction of coal and the aromatic fractions originating from the pyrolysis or gasification of coal and from shale oils or products derived from shale oils, alone or mixed, pretreated or not. Since the hydrogen is produced internally during the hydroreforming stage of the feed comprising at least one bio-oil, the hydrogen consumption for the co-treatment mixed with the hydrocarbon feedstocks derived from petroleum and / or coal is reduced.

 Preferably, said feed comprising at least one non-pretreated bio-oil may advantageously be a mixture of any of the abovementioned fillers, said feeds possibly being pretreated or not.

 hydroreforming

 According to the invention, said feedstock comprising at least one bio-oil, optionally at least one hydrocarbon feedstock derived from petroleum and / or coal, optionally other compounds derived from biomass, mixed with at least a part of the feedstock. organic phase of the effluent from the hydroreforming step which is recycled with a recycle rate equal to the mass flow rate of said organic phase on the mass flow rate of the non-pretreated bio-oil of between 0.05 and 2; , is introduced into said first hydroreforming step.

 The reaction is referred to herein as "hydroreforming" since it brings about a characteristic change in the molecular weight distribution of the feed comprising at least one bio-oil. Preferably, the hydroreforming step is considered as a hydrotreatment step in which the bio-oil is partially deoxygenated and partially hydroconverted with a partial production of hydrogen internal to the reaction.

 During said first hydroreforming step, the bio-oil is converted in order to stabilize the product, to make it miscible with the hydrocarbons, to cause an easy phase separation of the water in the bio-oil, d lowering its viscosity, lowering its corrosivity, significantly converting high molecular weight fractions to smaller molecules and lowering its oxygen content from about 50% to less than 15% by weight.

 The product of the hydroreforming stage

According to the invention, said first hydroreforming stage makes it possible to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase.

A gaseous phase is also advantageously obtained. Said gaseous phase containing predominantly C0 ₂ , H ₂ S and light C 1 -C 4 gases such as, for example, methane. Said gaseous phase may advantageously be separated from said aqueous and organic phases in any separator known to those skilled in the art.

 Said aqueous phase and said organic phase are advantageously separated by decantation.

 The analysis of the liquid effluent shows no sign of polymerization.

Said aqueous phase comprises water and may also comprise organic materials, mainly acetic acid and methanol (case of fast pyrolysis bio-oil), dissolved in said aqueous phase. The aqueous phase contains essentially water formed by hydro-deoxygenation and water contained in the unpreutered bio-oil which has not been converted and in general less than about 20% by weight of dissolved organic matter. For pyrolysis bio-oils, the aqueous phase separated in the first reforming step typically contains about 10% by weight of acetic acid and a lower proportion of methanol. Preferably, acetic acid and methanol are recovered in the aqueous phase because they are valuable byproducts. Acetic acid can advantageously be recovered by various known means such as distillation or evaporation, crystallization in the form of salt, for example an alkaline earth, or by solvent extraction, by means of ion exchangers. liquids for example. Because of its low boiling point and the absence of azeotropic formation with water, the methanol is simply recovered by distillation. Preferably, acetic acid and methanol taken together comprise at least 80% by weight of the organic components in the aqueous phase.

 The organic phase contains partially deoxygenated bio-oil and hydrocarbons. The term "partially deoxygenated bio-oil" means a bio-oil containing less than 25% by weight of oxygen and preferably less than 15% by weight of oxygen relative to the total weight of said organic phase. The organic phase therefore comprises less than 25% by weight of oxygen, preferably less than 15% by weight of oxygen, preferably less than 10% by weight of oxygen relative to the total weight of said organic phase.

 Thus, said organic phase is deoxygenated to a sufficient degree to make it miscible with hydrocarbon feeds such as those used in petroleum refining at high concentrations. In general, said organic phase can mix freely with most hydrocarbon products at concentrations of up to at least 50% by weight. Moreover, since the oxygen content is substantially reduced and the bio-oil is already partially converted, the hydrogen requirements during hydrotreatment / hydrocracking stages in the refinery are greatly minimized. In addition, said organic phase is stabilized so that it can be hydrotreated / hydrocracked in a refinery without the risk of large formation of solids.

 The organic phase preferably contains less than 25% by weight of oxygen and preferably less than 15% by weight of oxygen and less than 5% by weight of water, preferably less than 2% by weight of water relative to the mass. total of said organic phase. The hydrogen consumption during said first hydroreforming step is preferably less than about 2% by weight of the bio-oil.

 Preferably, the total acid number (TAN) is less than about 100 KOH / g oil. The total acid number is expressed in mg KOH / g of oil. This is the amount of potassium hydroxide in milligrams needed to neutralize acids in one gram of oil. There are standard methods for determining the acid number, such as ASTM D 974 and DIN 51558 (for mineral oils, biodiesels), or methods specific to biodiesels using European standard EN 14104 and ASTM D664, widely used worldwide.

The higher heating value (PCS) is greater than about 35 MJ / kg in the organic phase resulting from said first hydroreforming step. Without being exhaustive, said organic phase may for example be used in refining processes such as hydrocracking, hydrotreatment or catalytic cracking feedstock or as an industrial fuel.

The organic phase can also be directly used as a fuel-grade hydrocarbon product for road, air or sea transport. In the case where said organic phase is directly used for maritime transport, the quality of the fuels is fairly or very different depending on the grades considered (distillary fuels DMX, DMA, DMZ, DMB or residual fuels type RMA, RMB, RMD, RME, RMG or RMK according to the ISO 8217 standard) and the specifications specific to this type of fuels mean that the second hydrocracking step is not generally necessary. On the other hand, a second hydrotreating step operating under mild conditions, that is to say at high VVH, and under moderate hydrogen pressure, can advantageously be implemented. Preferably, said hydrotreatment step advantageously operates at a temperature of between 300 and 380 ° C., at a pressure of between 3 and 8 MPa and at a VVH between 1 and 3 h ^-1 . In the case of hydroreforming, the yields in the organic phase relative to the dry biomass, that is to say comprising 0% moisture, introduced into the pyrolysis step, are between 15 and 45% by weight, and preferably between 15 and 45% by weight. and 40% by weight The yields in the aqueous phase relative to the dry biomass are between 10 and 50% by weight, and preferably between 10 and 40% by weight.The gas phase yields relative to the dry biomass are between 5 and 30% by weight The conversion of the organic fraction of the bio-oil into partially deoxygenated bio-oil represents at least 70% by weight.

 We have also discovered that the partially deoxygenated bio-oil and more particularly the organic phase comprising at least said partially deoxygenated bio-oil produced by hydroreforming can easily be converted into fully deoxygenated hydrocarbons, compatible with the fuel pools, by means of modified processes. implemented in a conventional refinery, such as hydrotreating or hydrocracking.

 According to the invention, at least a portion of the organic phase of the effluent from the first hydroreforming step is recycled in said first hydroreforming step with a recycle ratio equal to the mass flow ratio of said organic phase. on the mass flow rate of the non-pretreated bio-oil of between 0.05 and 2 and in which the hydrocarbon effluent resulting from the hydrotreatment and / or hydrocracking step is not recycled in said first step of hydroreforming.

 Preferably, said recycle ratio is between 0.05 and 1.5, preferably between 0.2 and 1.3, and very preferably between 0.3 and 1.

The recycle rates used make it possible, overall, to have a sufficiently thermally stable charge to be able to be introduced into the preheating furnace as well as into the hydrorefaction reactor without risk of clogging, to facilitate the conversion and conversion reactions. deoxygenation of the bio-oil on the catalyst, to improve the dissolution of the gaseous hydrogen in the recycled organic phase resulting from the first hydroreforming step and to better manage the exothermicity generated by the reactions, while minimizing the recycling all or part of the organic phase resulting from hydroreforming in order to minimize the investment costs as well as the operating costs.

 The organic phase recycled in said first hydroreforming stage, and resulting from this stage, is also used to lower the density of the partially deoxygenated bio-oil produced during this first stage, which facilitates the phase clearance of the bio-oil. partially deoxygenated oil and the aqueous phase at the hydroreforming outlet. This separation which occurs easily during the cooling of the reaction mixture is another advantage that provides the presence of the recycled organic phase in the reaction mixture of the first hydroreforming step.

 Another advantage, the presence of the organic phase recycled in the reaction mixture tends to expel water from said organic phase, the water content is thus reduced as much as possible, which is favorable to the use of said organic phase as fuel or as a feedstock for subsequent hydrotreatment / hydrocracking steps. Thus, said phase said organic phase from the first hydroreforming step contains less than 5% by weight of water, preferably less than 2%.

 According to the invention, the catalyst used in the first hydroreforming step preferably comprises at least one transition metal chosen from the elements of groups 3 to 12 of the periodic table and at least one support chosen from activated carbons. , silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, alone or as a mixture, the catalyst preferably comprises a group 10 metal, alone or in combination with at least one metal from groups 3 to 12 and preferably at least one metal from groups 5, 6, 8, 9, 10, 1 1 of the classification periodic. Preferably, the catalyst comprises Ni, alone or in combination with at least one metal selected from Cr, Mo, W, Fe, Co and Cu.

 The metal content of the group 10 and preferably nickel, expressed as a percentage by weight of metal oxide of the group 10 is advantageously between 1 and 20% by weight, preferably between 5 and 15% by weight relative to the total mass of said catalyst.

 In the case where said metal of group 10 is used in combination with at least one metal of groups 3 to 12, the metal content of groups 3 to 12 is advantageously between 1 and 20% by weight relative to the total mass of said catalyst .

Preferably, the support is selected from silicas, transition aluminas, silica aluminas, and transition metal aluminates. These supports can be taken alone or in mixture. Preferred catalysts are Ni, NiCr or NiMn on porous carbon or NiMo on alumina or nickel aluminate.

 Preferably, said catalyst is prepared by conventional methods such as co-kneading or impregnation followed by one or more heat treatments.

 Said catalyst is advantageously used in the first hydroreforming step in reduced or sulphurized form.

 Operating conditions

Preferably, said first hydroreforming stage is carried out at a temperature of between 250 and 450 ° C., preferably between 250 and 400 ° C., and preferably between 270 and 360 ° C., and at an absolute pressure of between 3.4. and 27.6 MPa (500 and 4000 psi), preferably between 3.4 and 20.7 MPa (500 and 3000 psi), and most preferably between 6.9 and 2.7 MPa (1000 and 3000 psi) and more preferably between 15 and 20.7 MPa and at a high space velocity relative to the bio oil, preferably greater than 0.2 h ^-1 , preferably between 0.5 h ^-1 and 5 h ¹ , and more preferably between 1 hr ^-1 and 5 hr ^-1 .

 Preferably, the feedstock comprising at least the non-pretreated bio-oil is not preheated, or only preheated to a temperature below 80 ° before being introduced into the first hydroreforming step. Indeed, prolonged heating or storage at high temperature can cause deterioration.

 There are no limitations on the apparatus for carrying out the process. This can be conducted in batch or continuous mode. Nevertheless, for large-scale industrial applications, it will be preferable to operate continuously.

 In a preferred embodiment, the hourly space velocity is greater than about 1 kg of bio-oil per kg of catalyst per hour.

The hydroreforming reaction can be carried out in any reactor which facilitates the efficient dispersion of the bio-oil in the reaction mixture. A simple continuous stirred reactor (CSTR) is suitable in batch mode. For continuous applications, downflow, moving bed, bubbling bed or slurry reactors will be more suitable. While keeping the catalyst in the reactor, the moving-bed and ebull-bed reactor technologies have the advantage of allowing easy catalyst replacement during continuous operations, which gives greater flexibility to the operation of the reactor. unit, increases the duty cycle and maintains an activity as a function of time almost constant. These technologies also accept charges containing a few solids or producing solids during reactions, which is not the case for fixed bed reactor technology, except for the use of reactive guards reactors, for example. The technology of the slurry reactor, according to English terminology, has the same advantages as the moving bed and ebullated bed reactor technologies and the additional advantage of operating with a fresh catalyst, therefore with maximum activity, but the disadvantage to complicate a lot the recovery of the catalyst that comes out reactor with effluents. Other reactors meeting the principles set forth herein are within the abilities of those skilled in the art of chemical reactor technology and are within the scope of the present invention.

 Hydrogen consumption

 During the first hydroreforming step, an internal reforming of the bio-oil into carbon oxides and hydrogen takes place with a relatively high yield of hydrocarbons and a low relative consumption of hydrogen. The consumption of hydrogen is therefore moderate during the first hydroreforming step and the internal production of hydrogen by hydroreforming makes it possible to operate the process with a very limited external hydrogen consumption.

Part of the hydrogen necessary for the reaction in the first hydroreforming stage is generated internally, a priori from the water present in the bio-oil, by hydroreforming a portion of the bio-oil, in such a way that a significant portion of the oxygen is released as carbon dioxide, and the net hydrogen requirement is much lower than in the case of the direct hydrodeoxygenation process. The other part of the hydrogen necessary for the reaction is of external origin. Typically, the hydrogen consumption during the first hydroreforming stage is less than about 2% by weight of the mass of the bio-oil introduced into the reactor, and the corresponding emission of C0 ₂ to produce it is much less in the case of hydrodeoxygenation (HDO) type processes, with a consumption of at least 4 or 5% by weight of external hydrogen (depending on the oxygen content of the charge).

 The hydrogen of external origin can advantageously be supplied from fossil resources, by gasification / partial oxidation or by steam reforming, by steam reforming the bio-oil itself (although this implies a loss of carbon energy efficiency), by steam reforming methane produced and light gas fractions and / or fractions comprising light oxygen compounds from the first hydroreforming step or the subsequent hydrotreatment / hydrocracking step . It is also possible to react the CO recovered with the water produced to obtain hydrogen by means of the known reaction of gas to water conversion by "water gas shift" according to the English terminology (WGS) . Thus, the overall process, from the biomass to the final hydrocarbon product, could be self-sufficient in hydrogen.

 The first hydroreforming step therefore has the following advantages:

 the bio-oil is stabilized so that it can then be refined without the risk of coke formation, it is deoxygenated to a sufficient degree to make it miscible with typical hydrotreatment feeds in a refinery resulting from petroleum cuts,

the oxygen content in the bio-oil is reduced, so that the hydrogen requirements during a hydrotreatment in a refinery are greatly reduced, at a level adapted to conventional hydrotreatment / hydrocracking processes for petroleum products; hydrogen is generated internally from the water present by hydroreforming a portion of the bio-oil, so that the majority of the oxygen is released preferentially in the form of carbon dioxide and that the net need for external hydrogen is limited

 - the product is essentially free of water, the residual water forming a separate phase - the corrosivity of the product is greatly reduced since most of the organic acids are converted or pass into the aqueous phase

 light carboxylic acids such as acetic acid present in the crude bio-oil can be recovered in the aqueous phase and can be recovered as valuable by-products.

 - only a moderate amount of gaseous hydrocarbons, mainly methane, is formed

- if necessary, this methane fraction can be reformed by the so-called Steam Methane Reforming SMR process to produce biohydrogen for hydroreforming or hydrotreatment / hydrocracking processes

- It is also possible to use the bio-coal from the pyrolysis step as a feedstock for the gasification reactor to produce synthesis gas (CO + H ₂ ). This synthesis gas can also be used to produce bio-hydrogen for the hydroreforming and / or hydrotreating / hydrocracking processes.

 - Thus, by implementing the SMR process and / or the gasification of bio-coal, the process of conversion and upgrading of the bio-oil can be self-sufficient in hydrogen, without the need for hydrogen input. fossil origin

 - The weight yield of hydrocarbons is higher than that obtained by direct hydro-deoxygenation

 the operating conditions are relatively moderate and the hourly space velocity is high (residence time shorter than conventional hydro-deoxygenation processes) in continuous mode. In addition to increasing hydrocarbon yield and reduced hydrogen requirements, the goal is to reduce both capital and operating costs for upgrading bio-oils to produce a renewable source of chemicals and fuels for the environment. transport from whole lignocellulosic biomass.

 Hydrotreatment and / or hydrocracking step

Said organic phase comprising at least said partially deoxygenated bio-oil of the effluent from the first hydroreforming step still containing a high oxygen content, it is therefore not a pure hydrocarbon and it therefore requires a further processing to convert it into a fuel-grade hydrocarbon product for road, air or sea transport. In the case of a valuation for maritime transport, only a step of hydrotreatment under mild conditions can advantageously be implemented. According to the invention, at least a portion of said organic phase comprising at least said partially deoxygenated bio-oil of the effluent from the first hydroreforming stage is subjected to a second hydrotreatment and / or hydrocracking stage. . Said second hydrotreatment and / or hydrocracking step advantageously operates in the presence of hydrogen and of a hydrotreatment and / or hydrocracking catalyst, at a temperature above 200 ° C., at a pressure of between 2 MPa and 25 MPa and at a space velocity of between 0.1 hr ^-1 and 20 hr ^-1 .

At least a portion of said organic phase of the effluent from the first hydroreforming step is optionally introduced into said second hydrotreatment and / or hydrocracking step in mixture with a charge of fossil origin. Said filler of fossil origin preferably comprises at least 20% by volume of compounds whose boiling point is greater than 250 ° C. and preferably at least 50% by volume of compounds whose boiling point is greater than 350 ° C. ° C. Said charge of fossil origin is advantageously chosen from atmospheric distillates, vacuum distillates, such as, for example, distillates obtained from the direct distillation of crude or conversion units such as fluid catalytic cracking units, coking units or viscosity reductions, feedstocks from base lubricating oil aromatics extraction units or solvent dewaxing of lubricating base oils, distillates from hydrotreatment processes and / or hydroconversion of residues from hydroconversion to a fixed bed or moving bed or to a bubbling bed or suspension bed of atmospheric residues and / or vacuum residues and / or deasphalted oils obtained from any solvent deasphalting process using light solvents and deasphalted oils, or any mixture of the aforementioned fillers. The list above is not exhaustive. The n-C7 insoluble content of the overall charge is generally less than 5000 ppm, preferably less than 1000 ppm and more preferably less than 200 ppm.

hydrotreating

In a preferred embodiment, said second step is a hydrotreatment step operating in the presence of a heterogeneous catalyst and of hydrogen at a temperature of between 250 and 450 Ό, preferably between 300 and 400Ό, preferably between 320 and 400Ό. 380Ό and at a pressure between 2 MPa and 25 MPa, preferably between 5 MPa and 20 MPa and at an hourly space velocity between 0.1 hr ^"1 to 20 ^{h" 1,} and preferably between 0.5 h ^{" 1} and 5 h ^-1 and with a total amount of hydrogen mixed with the feedstock (including the chemical consumption and the amount recycled) such that the hydrogen / hydrocarbon volume ratio is between 100 and 3000 Nm ³ / m ³ and preferably between 100 and 1000 Nm ³ / m ³ .

Said hydrotreatment catalyst advantageously comprises an active phase comprising at least one Group 6 metal chosen from molybdenum and tungsten alone or as a mixture, preferably associated with at least one metal from groups 8 to 10, chosen from nickel and nickel. cobalt, alone or in admixture, and a support selected from the group consisting of alumina, silica, silica-alumina, magnesia, clays and mixtures of two or more of these minerals. Said support can furthermore contain other compounds such as, for example, oxides selected from the group consisting of boron oxide, zirconia, titania, phosphoric anhydride. Preferably, said support is made of alumina, and more preferably alumina [eta], [theta], [delta] or [gamma].

The total quantity of metal oxides of group 6 and groups 8 to 10 in the catalyst used is advantageously between 5 and 40% by weight, preferably between 6 and 30% by weight relative to the total mass of catalyst. In the case where said catalyst comprises nickel, the nickel oxide content is advantageously between 0.5 and 12% by weight and preferably between 1 and 10% by weight. In the case where the catalyst comprises molybdenum, the content of molybdenum oxide is advantageously between 1 and 35% by weight of molybdenum oxide (MoO ₃ ) and preferably between 5 and 30% by weight of molybdenum oxide, these percentages being expressed in% by weight relative to the total mass of catalyst. Said catalyst used in the hydrotreatment step of the method according to the invention must be advantageously characterized by a high hydrogenating power so as to lead as much as possible the selectivity of the reaction to a hydrogenation which retains the number of carbon atoms chains, in order to maximize the yield of hydrocarbons in the distillation range of middle distillates. Maximizing the hydrogenating function also makes it possible to limit the polymerization and / or condensation reactions leading to the formation of coke which may degrade the stability of the catalytic performance. A NiMo, NiW or CoMo catalyst will preferably be used.

 Said catalyst used in the second hydrotreating step may also advantageously contain at least one doping element chosen from phosphorus and / or boron. Said element may advantageously be introduced into the matrix or, preferably, deposited on the support. It is also possible to deposit silicon on the support, alone or with phosphorus and / or boron. The oxide content in said element is advantageously less than 20%, preferably less than 10% relative to the total mass of catalyst.

 The metals of the catalysts used in the second hydrotreating step of the method according to the invention may be metals or sulphurized metal phases. For maximum efficiency, these metal oxide catalysts are usually converted at least partially to metal sulfides. The catalysts based on metal oxides can be sulphurated by any technique known in the prior art, for example in the reactor (in situ) or ex situ, by contacting the catalyst, at elevated temperature, with a source. of hydrogen sulfide, such as dimethyl disulfide (DMDS).

Since biomass normally contains only a very small amount of sulfur, the use of non-sulphurized catalysts should avoid any risk of sulfur contamination in the fuels produced. Among the other catalysts based on metal oxides adapted to the hydrotreatment stage, mention may be made of the metal phases obtained by reduction in hydrogen. The reduction is generally carried out at temperatures of from about 150 ° to about 650 ° C, at a hydrogen pressure of from about 0.1 to about 25 MPa (14.5-3625 psi). Hydrocracking: mild hydrocracking and high pressure hydrocracking

In another preferred embodiment, said second step is a hydrocracking step advantageously operating at a pressure of between 2 MPa and 25 MPa, at a temperature of between 250 and 480 ° C., at a space velocity of between 0, 1 h ^"1 and 20 h ^{" 1} , and with a total amount of hydrogen mixed with the feed (including the chemical consumption and the amount recycled) such that the hydrogen / hydrocarbon volume ratio is between 80 and 5000 Nm ³ / m ³ . Preferably, at least a portion of said organic phase of the effluent from the first hydroreforming step is subjected to a second hydrocracking step in the presence of a heterogeneous catalyst and hydrogen.

 Said second hydrocracking step may advantageously be carried out in one or two stages according to the methods known to those skilled in the art.

 According to a preferred embodiment, said second hydrocracking step is a mild hydrocracking. In this case, said second hydrocracking step gives a net conversion of products whose boiling point is less than 375 ° is advantageously between 5 and 50% by weight, advantageously between 10 and 45% by weight.

Said second mild hydrocracking stage advantageously operates at a pressure of between 2 MPa and 12 MPa (290-1740 psi), preferably between 2 and 10 MPa (290-1450 psi), more preferably between 4 and 9 MPa (580 psig). -1305 psi) and more preferably between 3 and 7 MPa (435-1015 psi), at a temperature between 350 and 450Ό, at a space velocity of between 0.2 h ^-1 and about 5 h ^-1 and with a total amount of hydrogen mixed with the feed (including the chemical consumption and the recycled amount) such that the hydrogen / hydrocarbon volume ratio is between 100 and 2000 Nm ³ / m ³ , preferably between 200 and 1500 Nm ³ / m ³ .

 According to a preferred embodiment, said second hydrocracking step is a hydrocracking said high pressure. In this case, the operating conditions used in the process according to the invention make it possible to achieve conversions of products whose boiling point is less than 340 ° C., preferably less than 370 ° C., greater than 10%. and more, preferably between 20 and 95%, more preferably between 50 and 95%.

Said second so-called high-pressure hydrocracking step advantageously operates in the presence of hydrogen at a temperature of between 250 and 480 ° C., preferably between 320 and 450 ° C., more preferably between 330 and 435 ° C., at a pressure of between 5 MPa and 20 MPa (725-2900 psi), the hourly space velocity being between 0.1 hr ^-1 and 6 hr ^-1 , preferably between 0.2 hr ^-1 and 3 hr ^-1 , and with a quantity of hydrogen introduced is such that the volume ratio hydrogen / hydrocarbon is between 100 and 2000 Nm ³ / m ³ .

According to a very preferred embodiment, said second step is a hydrotreating step followed by a hydrocracking step of at least a portion of the organic phase of the effluent from the first hydroreforming step. The hydrotreating step upstream of the hydrocracking step makes it possible to reduce the heteroatoms potentially harmful for the hydrocracking catalyst (s). The hydrotreatment test step carried out in the presence of a conventional hydrorefining catalyst known to those skilled in the art.

 Said second hydrotreatment step and / or hydrocracking may advantageously be carried out in a fixed bed reactor or in a so-called driven bed: moving bed reactor, bubbling bed reactor or suspension reactor. It is possible to use a single catalyst or several different catalysts, simultaneously or successively, in the case of a fixed bed reactor. Said steps can be carried out industrially in one or more reactors, with one or more catalytic beds. The reaction exotherm during the hydrotreatment is limited by any method known to those skilled in the art: recycling of the liquid product, cooling with recycled hydrogen, etc.

 Said second hydrocracking stage may be carried out in one or more fixed or entrained catalyst beds or in a fixed and driven bed, in one or more reactors, in the context of a one or two stage hydrocracking scheme as it is described in US-2007 / 0.209.968, with or without liquid recycling of the unconverted fraction, optionally in combination with a hydrorefining catalyst upstream of the catalyst of the present invention.

The hydrocracking catalysts used in the second hydrocracking step are all of the bifunctional type which combines an acid function and a hydro-dehydrogenating function. The acid function is provided by substrates which have surface areas generally of between 150 and 800 m ² / g and a surface acidity, such as halogenated alumina (in particular chlorinated or fluorinated alumina), aluminas, zeolites and silica-aluminas. amorphous, or their combinations. The hydro-dehydrogenating function is provided either by one or more metals from Groups 8 to 10 of the Periodic Table of Elements, or by a combination of at least one Group 6 metal and at least one Group 8 metal to The catalyst may comprise metals of Groups 8 to 10, for example nickel and / or cobalt, most often in combination with at least one Group 6 metal, molybdenum and / or tungsten, for example. It is possible to use, for example, a catalyst comprising between 0.5 and 10% by weight of nickel (expressed in terms of nickel oxide NiO) and between 1 and 30% by weight of molybdenum, preferably between 5 and 25% by weight. of molybdenum (expressed in terms of MoO ₃ molybdenum oxide) on an amorphous mineral substrate. The total content of the Group 6 and Group 8 to 10 metal oxide catalyst is generally between 5 and 40% by weight, preferably between 7 and 30% by weight. In the case where the catalyst comprises at least one group 6 metal associated with at least one non-noble metal of groups 8 to 10, the catalyst is preferably a sulphurized catalyst.

The catalyst may also contain a promoter element such as phosphorus and / or boron. This element may have been introduced into the matrix or, preferably, deposited on the substrate. Silicon may also have been deposited on the substrate, alone or with phosphorus and / or boron. The catalysts preferably contain silicon deposited on a substrate such as alumina, optionally with phosphorus and / or boron deposited on the substrate and also containing minus a Group 8 to 10 metal (Ni, Co) and at least one Group 6 metal (Mo, W). The concentration of said element is usually less than about 20% by weight (oxide base) and most often less than about 10% by weight. The concentration of boron trioxide (B ₂ O ₃ ) is usually between 0 and 10% by weight.

 Other conventional catalysts comprise an FAU type zeolite Y, an amorphous support

(generally alumina) and at least one metal or metal compound having a hydro-dehydrogenating function (generally at least one Group 6 metal and 8 to 10 groups, preferably at least one Group 6 metal and at least one non-noble metal groups 8 to 10).

 Another example of suitable catalysts are the so-called composite catalysts described in document US-2007 / 0.209.968, which comprise at least one metal or a metal compound having a hydro-dehydrogenating function chosen from group 6 and groups 8 to 10. , a support based on silica-alumina and at least one zeolitic support.

 The hydrocracking catalyst preferably comprises at least one metal of groups 8 to 10 chosen from Ni and / or Co, and at least one metal of group 6 chosen from Mo and / or W, on a silica support -alumine and / or zeolite.

 Before the injection of the feedstock, the catalysts employed in the process according to the present invention are preferably subjected to a preliminary sulphurization treatment in order to convert, at least in part, the oxide species into sulphides before they are brought into contact with the load to be treated. The sulfidation activation treatment is known to those skilled in the art and can be carried out using any method already described in the literature, either in situ, that is to say in the reactor, or ex situ.

 The hydrocarbon effluent obtained at the end of the second hydrotreatment and / or hydrocracking step is a completely deoxygenated liquid hydrocarbon effluent. Said hydrocarbon liquid effluent comprises a content of hydrocarbon compounds advantageously greater than 98% by weight and preferably greater than 98.9% by weight. Said effluent also advantageously comprises less than 1% by weight of oxygen and less than 0.1% by weight of water relative to the total mass of said effluent.

 The higher heating value (PCS) of said hydrocarbon liquid effluent is about 45 MJ / kg, which is comparable to that of diesel. The boiling point of said hydrocarbon liquid effluent, measured by simulated distillation by gas chromatography, according to the ASTM method D2887 for example, is less than about 500 ° C, preferably less than 450 ° C.

The overall yield of hydrocarbon liquid effluent obtained at the end of the second hydrotreatment and / or hydrocracking step relative to the dry biomass is advantageously between 10 and 48% by weight and preferably between 10 and 40% by weight. The hydrogen consumption during the second stage is preferably less than 2% by weight relative to the mass of bio-oil introduced. This is equivalent to a hydrogen consumption during the second stage is preferably less than 4% by weight relative to the mass of organic effluent sent to said second stage. Said hydrocarbon liquid effluent from the second hydrotreatment and / or hydrocracking step advantageously undergoes a distillation step to obtain fuel base fractions that can be incorporated in the diesel fuel or kerosene pool.

 Figure 1 illustrates the method according to the invention.

The bio-oil resulting from the rapid pyrolysis of the biomass is introduced via line (1), mixed with hydrogen (3) and the organic phase (8) resulting from the hydroreforming step in the first zone. hydroreforming (2). The liquid effluent from the hydroreforming zone via line (4) is introduced into a separator (5). A gaseous phase predominantly containing CO 2, H ² S and light C 1 -C 4 gases such as, for example, methane is separated via line (7) and an aqueous phase is also separated via line (6). An organic phase containing the partially deoxygenated bio-oil and hydrocarbons is withdrawn through line (8) and a portion of said organic phase is recycled to the hydroreforming zone. The non-recycled portion of said organic phase is then sent to a second hydrotreatment and / or hydrocracking zone (9). The hydrocarbon liquid effluent resulting from the second hydrotreatment and / or hydrocracking step (10) then undergoes a distillation step to obtain at least the fuel base sections (12), (13), (14) that can be incorporated in the diesel fuel pool, kerosene.

 The examples below illustrate the invention without limiting its scope.

 The experiments were conducted with a bio-oil obtained by rapid pyrolysis of hardwood mixtures. The operating conditions of the rapid pyrolysis are a temperature of 500 ° C. and a residence time of the vapors of the order of one second. Table 1 below presents the data obtained following the analysis of this bio-oil.

 Table 1: Analysis of bio-oil

Density at 20 ° C, g / cm ³ 1, 21

Cynematous viscosity at 20 ° C, mm ² / s 182.6

PH 2.5

Higher heating value, MJ / kg ^at 18.03

Lower heating value, MJ / kg ^b 16,46

Water content,% mass 20.7

Pyrolytic lignin,% mass 19.0

Carbon, mass% 43.9

Hydrogen, mass% 7.39 Nitrogen,% mass <0.05

Oxygen, mass% 47.3

Sulfur (ppm) 92 ^a The higher calorific value is measured with a calorimetric bomb, ^b the lower calorific value is calculated using the following equation: LHV (J / g) = HHV (J / g) - 218.13 ^* H% (% wt)

Example 1 Preparation of C1 and C2 Hydroreforming Catalysts

Catalyst C1 is a NiCr / C formulation catalyst. The precursors of nickel and chromium, in their nitrate form (respectively Ni (NO ₃ ) ₂ and Cr (NO ₃ ) ₃ ), are provided by AIdrich. The support is an active carbon in cylindrical extruded form supplied by the company Norit (RX3 Extra). This catalyst C1 is obtained by dry impregnation of an aqueous solution containing the precursors of nickel and chromium. The volume of solution is equal to the volume of water uptake of the support (that is to say the maximum volume of water that can penetrate its porosity). The concentrations of nickel and chromium precursors in solution are determined so as to obtain the target levels on the final catalyst: 10% by weight of nickel and 5% by weight of chromium. After impregnation of this aqueous solution, the catalyst is allowed to mature at room temperature for 4 hours in a chamber saturated with water, then dried in an oven at 70 ° C. under air for 3 hours and finally subjected to a heat treatment at 300 ° C. for 3 hours under nitrogen (at a flow rate of 1.5 L / h / g of catalyst).

Catalyst C2 is of ΝίΜο / δ-alumina formulation. The support is a transition alumina of delta structure and is in the form of a trilobal extrusion. This support is obtained by dehydration of boehmite in transition alumina. Beforehand, the boehmite powder undergoes a kneading / extrusion step in order to obtain boehmite shaped into an extruded three-lobed form. The δ-alumina support of the catalyst is finally obtained after calcination at 900 ° C. in air of the formed boehmite. The target levels for this catalyst are respectively 6% by weight for nickel and 5% by weight for molybdenum. The synthesis protocol consists in carrying out a dry impregnation of a solution of ammonium heptamolybdate and nickel nitrate. The concentrations of the precursors in the solution are adjusted so as to deposit on the support the weight contents of metals. The solid is then allowed to mature at room temperature for 12 hours in a chamber saturated with water, and then dried in an oven at 120 ° C in air for 3 hours. Finally, the solid is calcined at 500 ° C. for 2 hours under an air stream (at a flow rate of 1 L / h / g of catalyst). Table 1. Formulation of catalysts C1 and C2.

EXAMPLE 2 Production of a Hydrocarbon Liquid from an Unprocessed Bio-Oil with a Recycle Ratio of 0.33 This example illustrates the production of an oxygen-free hydrocarbon liquid from a biocharged charge. oil not pretreated with the characteristics described in Table 1 mixed with an organic liquid phase from the first hydroreforming step at a ratio of 1: 3 (organic liquid phase: bio-oil) simulating a recycle rate of 0.33 . The feed containing the unprocessed bio-oil is first introduced into an autoclave reactor with a hydroreforming catalyst for the first hydroreforming step.

15 g of the catalyst C1 prepared according to Example 1 are reduced in a reduction cell at 300 ° C. for 3 hours under 30 NL / h of hydrogen and then introduced, in an inert glove bag, into a basket which is placed in a the autoclave reactor with 75 g of bio-oil and 25 g of the organic liquid phase from the hydroreforming step. The reactor is sealed and purged with nitrogen and then with hydrogen. The reactor is then pressurized to 4.83 MPa (700 psia) with 18.37 LH ₂ , then gradually heated with stirring to 330 Ό and maintained at this temperature for 3 hours during which the maximum pressure reached was 13 ° C. , 9 MPa (2016 psia). The equivalent space velocity, defined as the ratio of the volume of bio-oil introduced on the product of the catalyst volume by the duration of the test, is equal to 0.7 h ^-1 .The reactor is then rapidly cooled and brought to temperature. The reaction gas phase, 4.9 NL in total, is sent to a collector bottle and analyzed by gas chromatography. it contains 63.9% of H ₂ , 21.6% of C0 ₂ , 10.6% of CH ₄ and 3.0% of CO by volume.

The basket containing the catalyst is recovered from the product liquid, which is separated into 50.33 g of a higher homogeneous organic liquid liquid phase of density 980 kg / m ³ and 38.37 g of a lower aqueous phase. No tar deposit is observed in the reactor. The organic liquid phase contains 13.93% by weight of elemental oxygen and 4.68% by weight of water (2.36 g) of water, determined by elemental analysis and Karl-Fischer assay. The overall net hydrogen consumption is 15.26 NL (1.39 g), which corresponds to about 1.85% by weight relative to the mass of bio-oil introduced. From the above data, organic liquid phase and aqueous phase yields are determined at 27.8% and 21.2% by weight, respectively, based on dry biomass containing 0% moisture.

The organic liquid phase resulting from the hydroreforming step is then sent to a fixed-bed reactor which is crossed in the presence of a conventional hydrotreatment catalyst for the second hydrotreating step (HDT). The hydrotreatment catalyst used comprises 4.3% by weight of NiO, 21% by weight of MoO ₃ and 5% by weight of P ₂ 0 ₅ supported on a gamma alumina. Prior to the test, this catalyst is sulphurized in situ at a temperature of 350 ° C., using a feedstock containing heptane additive of 2% by weight of dimethyl disulphide (DMDS).

25 ml / h of the organic liquid phase resulting from the hydroreforming step is introduced into an isothermal reactor and fixed bed charged with 25 ml of hydrotreatment catalyst. The corresponding space velocity is equal to 1 hr ^-1 800 Nm ³ of hydrogen / m ³ of feed are introduced into the reactor maintained at a temperature of 320 Ό and a pressure of 10 MPa (1450 psia). In the sulphide catalyst, 50 ppm by weight of sulfur in the form of DMDS is added to the feedstock Under the reaction conditions, the DMDS is completely decomposed to form methane and H ₂ S.

The test is carried out for 1 hour and made it possible to recover 19.63 g of a hydrocarbon liquid phase with a density of 842 kg / m ³ , containing 0.12% by weight of water and 0.2% by weight of elemental oxygen ( analyzer limit of quantification) and 99.7% by weight of hydrocarbons. The net hydrogen consumption corresponds to 2.2% by weight relative to the mass of organic effluent sent to the hydrotreatment stage, which corresponds to a net hydrogen consumption of 0.9% by weight relative to the bio-oil. The overall yield of hydrocarbon liquid is equal to 21.4% by weight on the basis of dry biomass.

 These results show that a totally deoxygenated hydrocarbon product according to the invention can be obtained by hydroreforming the bio-oil with a recycling rate of the organic effluent resulting from the hydroreforming step of 0.33 and the catalyst C1, followed by hydrotreating of the organic effluent from the hydroreforming step without recycling the hydrocarbon product from the hydrotreating step in the hydroreforming step.

 Example 3 Production of a Hydrocarbon Liquid from an Unprocessed Bio-Oil with a Recycling Rate of 1

This example uses the conditions of Example 2 except that the unpretreated bio-oil is introduced in admixture with the organic liquid phase resulting from the hydroreforming stage with a 1: 1 ratio simulating a recycle rate of 1. and the hydroreforming catalyst is different.

15 g of the catalyst C2 prepared according to example 1 are reduced in a reduction cell at 400 ° C. for 3 h under 30 NL / h of hydrogen and then introduced, in an inert glove bag, into a basket which is placed in the autoclave reactor with 50 g of bio-oil and 50 g of the organic liquid phase from the hydroreforming step. The autoclave is closed tightly, put under pressure at 6.7 MPa (971 psia) with 25.43 ° NL of H ₂ , then gradually heated with stirring at 330 ° C., maintained at this temperature for 3 hours during which the maximum pressure reached was 1. 7.3 MPa ( 2509 psia). The equivalent space velocity, defined as the ratio of the volume of bio-oil introduced on the product of the catalyst volume by the duration of the test, is equal to 0.8 h ^-1 .The reactor is then rapidly cooled and brought to temperature. The reaction gas phase, totaling 18.63 L PTN, is removed and analyzed, and gas chromatographic analysis indicates that the reaction gas phase is at a pressure of 3.1 MPa (449 psia). 85.9% H ₂ , 6.9% C0 ₂ , 0.9% CO, 2.2% CH ₄ and 3.1% C2 + by volume.

The basket containing the catalyst is recovered from the product liquid, which is separated into 70.43 g of a higher homogeneous organic liquid liquid phase of density 950 kg / m ³ and 22.78 g of a lower aqueous phase. No tar deposit is observed in the reactor. The organic liquid phase contains 9.74% by weight of elemental oxygen and 0.98% by weight of water (0.84 g) of water, determined by elemental analysis and Karl-Fischer assay.

 The overall net hydrogen consumption is 9.45 NL (0.844 g), which corresponds to about 1.69% by weight based on the bio-oil that has been reacted. From the above data, organic liquid phase and aqueous phase yields are determined at 31.6% and 10.2% by weight, respectively, on the basis of dry biomass containing 0% moisture.

 The second hydrotreating step of the organic liquid phase resulting from the hydroreforming step is carried out according to the operating conditions described in Example 3. The hydrotreatment catalyst of Example 3 is used for this example. It is preliminary to the sulphuretted test in situ at a temperature of 350 ° C., using a heptane feed additive of 2% by weight of dimethyl disulphide (DMDS). 50 ppm by weight of sulfur in the form of DMDS is also added to the feed to maintain the catalyst in the sulfide state.

21.34 g of a hydrocarbon liquid phase with a density of 830 kg / m ³ , containing 0.01% by weight of water, 0.2% by weight of elemental oxygen (limit of quantification of the analyzer) and 99.8 % mass of hydrocarbons is recovered after 1 hour of operation of the test. The net consumption of hydrogen corresponds to 1.8% by weight relative to the mass of organic effluent sent to the hydrotreatment stage, which corresponds to a net hydrogen consumption of 0.8% by weight relative to the bio-oil. The overall yield of hydrocarbon liquid is equal to 25.6% by weight on the basis of dry biomass.

 These results show that a completely deoxygenated hydrocarbon product according to the invention can be obtained by hydroreforming the bio-oil with a recycling rate of the organic effluent from the hydroreforming stage of 1 and a catalyst C2 followed hydrotreating the organic effluent from the hydroreforming step without recycling the hydrocarbon product from the hydrotreating step in the hydroreforming step.

Claims

1. A process for producing hydrocarbon products from a feedstock comprising at least one non-pretreated bio-oil, said feedstock being introduced in at least the following steps:

 a first hydroreforming step in the presence of hydrogen and a hydroreforming catalyst comprising at least one transition metal chosen from the elements of groups 3 to 12 of the periodic table and at least one support chosen from activated carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, taken alone or as a mixture, to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase,

 a second step of hydrotreating and / or hydrocracking at least a portion of the organic phase of the effluent from the first hydroreforming step to obtain a hydrocarbon effluent,

 wherein at least a portion of the organic phase of the effluent from the first hydroreforming step is recycled to said first hydroreforming step, in admixture with said feed comprising at least one bio-oil, with a recycle rate equal to the ratio of the mass flow rate of said organic phase to the mass flow rate of the non-pretreated bio-oil of between 0.05 and 2 and in which the hydrocarbon effluent resulting from the hydrotreating and / or hydrocracking step is not recycled in said first hydroreforming step.

2. Method according to claim 1, wherein said recycling rate is between 0.05 and 1.5,

3. Method according to claim 2, wherein said recycling rate is between 0.2 and 1.3.

4. The method of claim 3, wherein said recycling rate is between 0.3 and 1.

5. Method according to one of claims 1 to 4, wherein the bio-oil is produced by rapid pyrolysis from a biomass feedstock.

6. Method according to one of claims 1 to 5, wherein said hydroreforming catalyst comprises Ni, alone or in combination with at least one metal selected from Cr, Mo, W, Fe, Co and Cu.

 7. Method according to one of claims 1 to 6, wherein the support of said hydroreforming catalyst is selected from silicas, transition aluminas, silica aluminas, and transition metal aluminates, alone or in admixture.

8. Process according to one of claims 1 to 5, wherein said hydroreforming catalyst is a Ni, NiCr or NiMn on porous carbon catalyst or NiMo on alumina or nickel aluminate.

9. Method according to one of claims 1 to 8, wherein said first hydroreforming step is carried out at a temperature between 250 ° C and 450 ° C at an absolute pressure between 3.4 and 27.6 MPa , at a space velocity exceeding 0.2 h ^"1 .

10. Method according to one of claims 1 to 9, wherein said second hydrotreating step and / or hydrocracking operates in the presence of hydrogen and a hydrotreating catalyst and / or hydrocracking, at a temperature above 200 Ό, at a pressure of between 2 MPa and 25 MPa and at a space velocity of between 0.1 hr ^-1 and 20 hr ^-1 .

1 1. The process according to claim 10, wherein at least a portion of said organic phase of the effluent from the hydroreforming step is introduced into said second hydrotreating and / or hydrocracking step in admixture with a feedstock. fossil origin.

12. Method according to one of claims 1 to 1 1, wherein said second step is a hydrotreating step operating in the presence of a heterogeneous catalyst and hydrogen at a temperature between 250 ° C and 450Ό, at a pressure between 2 MPa and 25 MPa, and at a space velocity of between 0.1 h ^"1 and 20 h ^{" 1} .

13. Method according to one of claims 1 to 1 1, wherein said second step is a hydrocracking step operating in the presence of a heterogeneous catalyst and hydrogen at a pressure of between 2 MPa and 25 MPa, at a temperature between 250 ° C and 480 ° C, at a space velocity of between 0.1 hr ^-1 and 20 hr ^-1 , and with a total amount of hydrogen mixed with the feedstock (such as hydrogen / hydrocarbon volume ratio is between 80 and 5000 Nm ³ / m ³ .

14. The method of claim 13, wherein said second step is a hydrotreating step followed by a hydrocracking step of at least a portion of the organic phase of the effluent from the first hydroreforming step.